Influence of Riser Length of a Fluid Catalytic Cracking Pilot Plant on

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THE INFLUENCE OF RISER LENGTH OF AN FCC PILOT PLANT ON CATALYST RESIDENCE TIME AND PRODUCT SELECTIVITY Angelos A. Lappas, Dimitrios K. Iatridis, Evie P. Kopalidou, and Iacovos A. Vasalos Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.7b00547 • Publication Date (Web): 19 Apr 2017 Downloaded from http://pubs.acs.org on April 25, 2017

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Industrial & Engineering Chemistry Research

THE INFLUENCE OF RISER LENGTH OF AN FCC PILOT PLANT ON CATALYST RESIDENCE TIME AND PRODUCT SELECTIVITY Angelos A. Lappas1*, Dimitrios K. Iatridis1, Evie P. Kopalidou1 and Iacovos A. Vasalos1 1

Chemical Process and Energy Resources Institute (CPERI), Centre for Research and Technology Hellas (CERTH) 6th km Harilaou-Thermi Road, 57001 Thermi, Thessaloniki, Greece *

[email protected]

Abstract The design and operation of a circulating Fluid Catalytic Cracking (FCC) riser pilot plant is presented in the paper. Emphasis is given on the effect of riser length on feedstock conversion and coke yield. Riser lengths considered in the study were 1.46 and 9.95 meters. The effect of catalyst residence time on feedstock conversion and coke yield at constant feed rate and reactor temperature (526°C) is presented. The catalyst residence time varies with reactor volume and catalyst circulation rate. It is shown that for the same feed conversion, high catalyst circulation rates are required with the lower riser volume. Lower gasoline and lower coke yields are experimentally observed with the short riser length. A second order kinetic model is used to correlate the variation in conversion and coke yield. Commercial results validate the pilot plant design as a useful tool for guiding commercial operations. Keywords: FCC, pilot plant, riser reactor, feedstock effects

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INTRODUCTION It is recognized that the profitability of a new or existing commercial process is greatly enhanced with the design and operation of a pilot plant. In the Fluid Catalytic Cracking (FCC) process a pilot plant makes possible the evaluation of new feedstocks and catalysts in advance and the selection of operating conditions for optimizing the product yields in a commercial unit. These facts have been reported by Corma et al.1, who in a recent article reviewed the available laboratory units for ranking cracking catalysts prior to their commercial use. These units include: Microactivity Catalyst Testing (MAT), Fixed Fluid Bed (FFB) and Transported Fluid Bed (TFB) reactors. Corma et al.1 presented data showing the catalyst deactivation taking place with coke deposition on the catalyst. The catalyst deactivation is the key variable influencing feed conversion and selectivity. Because a circulating catalyst riser pilot plant with continuous catalyst regeneration results in coke deposition and uses operating conditions similar to a commercial unit, the obtained product yields in a pilot plant match those of a commercial unit. The above information has been verified in an article by Bryden et al.2, where extensive data show comparative results between the Grace Davison Circulating Riser (DCR) TFB pilot plant and a commercial unit. It was shown that the pilot plant matches with great accuracy the operating conditions and product yields of a commercial unit. The same article also establishes the usefulness of a pilot plant in testing unconventional feedstocks like bio-oils and straight run shale oil. The validation of pilot plant with commercial data has been the subject of a presentation by Joyal3 in the National Petroleum Refinery Association meeting in San Diego in 2008. Using pilot plant data from the CPERI pilot plant4 and commercial audited data from several BP units,


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it was concluded that the evaluation of cracking catalysts in advance delivered significant value of several million dollars per year to the refineries. In a very important study, Alvarez-Castro et al.5 have also reported a comparison of commercial data with model predictions. The predictions were based on a rigorous set of equations describing the hydrodynamics of the riser with CFD tools, used in tandem with a 12lump kinetic model developed by Wu et al.6 These new methods permit the examination of the effect of unit operating conditions on the hydrodynamic features of the riser, feed conversion and product selectivity. It was reported that when the feed and catalyst temperature at the reactor inlet is fixed, then the catalyst to oil ratio has a major effect on the yield distribution. Bollas et al.7 reported results from a short length riser (1.46 m) FCC pilot plant and a comparison with a model, based on empirical and fundamental correlations and combining hydrodynamic and kinetic theories of fluid catalytic cracking. The model takes into account the specific riser geometry and a second order kinetic scheme describing the feedstock conversion as a function of catalyst residence time in the reaction zone. Excellent agreement was obtained among model predictions and experimental observations such as riser hydrodynamic parameters, feedstock conversion and coke yield. Lappas et al.8 also reported comparative results for gas oil cracking and an equilibrium cracking catalyst for two different units: a confined fluid bed bench scale unit and a pilot plant transported fluid bed riser reactor with continuous catalyst regeneration4. It was reported that although both units correctly rank the feedstocks and catalysts, compared with commercial units, there are significant differences in the absolute yields obtained at constant feed conversion.



Let us now review the factors, which according to literature, are considered important in the design and performance of the reactor section of a Fluid Catalytic Cracking Unit (FCCU) pilot plant: Feed atomization: Atomization of the feed is the first step in the sequence of events which leads to catalytic cracking. In a commercial unit the feed is injected near the bottom of the riser with multiple nozzles9 in order to achieve the feed distribution in small droplets and thus achieve fast feed vaporization. This is a requirement to avoid thermal cracking reactions in the feed contact with the high temperature regenerated catalyst. In a pilot plant reactor, because of the small diameter of the riser entrance, a single nozzle is used. An atomizing nozzle consists of a small orifice used to convert a continuous feed stream into fine droplets with the use of a gaseous stream, usually steam. For practical reasons steam is replaced with nitrogen in laboratory studies. There has been an evolution in the nozzle design over the years. The most commonly nozzle geometry involves mixing of a gaseous and a liquid stream upstream of a small orifice10. Although the mechanism of the atomization is not well understood, several empirical correlations exist10, which describe the effect of the physical properties of the liquid and the effect of the relative gas and liquid rates on the size of the droplet formed at the nozzle exit. A big step was taken in the nozzle design with the development11 of Effervescent atomization in combustor design. In Effervescent atomization a small amount of gas is injected inside the atomized body and this leads to significant improvements in performance in terms of small droplet size. This results in lower pressure drops and lower exiting velocities of the gasliquid stream. Because the momentum issued from a nozzle operating with effervescent


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atomization is lower than a single liquid stream, the jet penetration in a fluidized bed is expected to be lower. Based on Effervescent atomization, alternative nozzle designs were recently described by Jedelsky et al.12,13. The geometry of the mixing chamber and the size of the exiting orifice were found to be critical parameters for the droplet size of the liquid spray. Vaporization of FCC feed in riser entrance: The oil droplets, exiting the nozzle, are vaporized either with a homogeneous or a heterogeneous vaporization. For homogeneous vaporization, as reported by Buchanan14 and Nguyen et al.15, the estimated vaporization time for an average droplet size of 100 µm is in the range of 15 to 25 ms. For heterogeneous vaporization and the same droplet size (100 µm), the estimated vaporization time is in the range of 0.3 to 4 ms. Hence, it seems that the estimated vaporization times for the homogeneous model are an order of magnitude higher than those for the heterogeneous model. The heterogeneous model is heat transfer driven and for this reason it results in lower vaporization times than the homogeneous model, which is mass transfer driven15. The heterogeneous model considers direct contact of the oil droplets with hot catalyst particles. Mirgain et al.16 have shown that the most effective contact of oil droplets with catalyst particles is achieved when the oil is injected into a catalyst bed of intermediate gas void fraction (0.70 to 0.95). As a result good mixing of all phases is achieved. Fluidization dynamics - Feed injection in the bottom of the riser: The small droplets emerge through the tip of a nozzle in the form of a jet at high velocities of about 50 m/s10. Jet Penetration is the second step in the sequence of events which leads to FCC feed cracking. Jet penetration can be defined from the momentum profiles as that height on the axis at which the profiles level out. Experimental work carried out by Behie et al.17,18 has established that the



momentum issuing from the jet dissipates within a cone of angle Θ. In a fluidized bed of the FCC, this angle was found equal to 36°. Solids and gases from the region near the wall of the column are drawn into the dissipation cone and contribute to the reduction of the momentum by slowing the jet down. The same conclusions were reached in a study by Chenglin et al.19, who investigated the effect of the ratio of the momentum of the jet gas and the lift superficial velocity on the height, where complete mixing takes place. It was shown that the height of the full mixing decreases with the decreasing momentum ratio. Fluidization - Riser inlet: The vaporized gas oil feed and steam are mixed with an FCC catalyst in the bottom of the riser. The vapor products together with the steam provide the gases for the fluidization of the catalyst particles and their rapid transfer out of the mixing zone into the riser. Calculation of the average solids holdup and the solids density (ρb) distribution in the entrance zone is one of the key requirements for either data correlations or detailed reactor model development. Avidan and Yerushalmi20 presented a generalized correlation, which is valid for gas-solids fluidized systems at gas superficial velocities exceeding 0.5 m/s. This correlation is often represented by the well-known Richardson-Zaki21 equation: Ug / Vt = εgn

(1)

where: Ug superficial gas velocity, m/s Vt particle terminal velocity, m/s εg void fraction n index dependent on the particle to tube diameter and the particle Reynolds number


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Bollas et al.7 have reported data on the calculation of void fraction in the entrance zone of a riser reactor with the same diameter as the riser used in this work. It has been shown that for gas superficial velocities from 0.15 to 0.25 m/s in the entrance zone the calculated void fraction is in the range of 0.69 to 0.82. Fluidization - Vertical Riser: In addition to the catalyst density in the riser inlet, another consideration is to estimate the catalyst density in the fully developed section of the riser. This is a function of the non slip catalyst density (ρNS) and the catalyst particle slip factor (ψ). The non slip catalyst density is the mass of catalyst per unit reactor volume, assuming that the catalyst particle and the gas velocity are the same (slip particle velocity is zero). The slip catalyst density in the riser section is calculated by multiplying the slip factor by the non slip catalyst density (ρNS), which is calculated by the following equation: ρNS = (Qc + Qv*ρv) / (Qc/ρp + Qv)

(2)

where: Qc catalyst flow rate, kg/h Qv volumetric flow rate, m3/h ρv vapor density, kg/m3 Qv*ρv vapor flow rate, kg/h ρp catalyst particle density, kg/m3 The slip factor is defined here as the ratio of the gas velocity to the particle velocity. The estimation of the slip factor (ψ) has been the subject of intense industrial research. For example, Matsen22 reported that in large commercial risers the slip factor ψ is approximately equal to 2, and hence, the particle velocity is Ug/2εg.



Patience et al.23 have presented the following correlation used for the hydrodynamically well developed region of the riser and for superficial gas velocity in the range of 6 to 12 m/s: ψ = 1 + 5.6/Fr + 0.47/Frt0.47

(3)

where: Fr Froude number, Ug/(gD)0.5 Frt Froude number, Vt/(gD)0.5 D riser diameter, m g gravitation constant, m/s2 Another method to estimate the catalyst density in the riser is from differential pressure drop measurements. Pugsley and Berruti24 have reported that the void fraction εg for the hydrodynamically fully developed region of a riser at gas superficial velocities less than 5 m/s, can be estimated by assuming that the time average pressure drop is attributable only to the hydrostatic head of solids εg = 1 - dP/(ρp g dz)

(4)

where: dP/dz axial pressure gradient, Pa/m Das et al.25 reported data for an air-FCC riser circulating catalyst system over a range of catalyst fluxes from 12.5 to 50 kg/m2/s and gas superficial velocities from 2.9 to 4.9 m/s. Using riser pressure data and Eq. (4), the calculated gas voidage was in the range from 0.90 to 0.9998. From the gas voidage, the catalyst density in the riser can be calculated. This is essential for estimating the catalyst holdup in the riser. The feed rate divided by the catalyst holdup defines the Weight Hourly Space Velocity (WHSV).


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Kinetics of gas oil conversion: In the FCC process, a gas oil feed mainly consisting of thousands of species with a boiling point range exceeding 221°C, is converted to gasoline, light hydrocarbons and coke. Feed volume or weight conversion is equivalent to 100 minus the volume or weight fraction of liquid products boiling over 221°C. For the refiner, it is important to be able to predict in advance the feed conversion and product selectivity as a function of feed quality, catalyst properties and unit operating conditions. When detailed feed characterization data are missing, a simplistic approach is to represent the complex feed with a single pseudo component, the weight fraction of the feed boiling over 221°C. Using this simplistic approach, Blanding26 and Weekman27 found that the gas oil conversion follows a second order reaction coupled with catalyst decay, according to the following equations: R(y, θ) = ko y2 e-λθ

(5)

λ = α tc

(6)

where: R(y, θ) instantaneous kinetic rate expression ko intrinsic reaction velocity constant at θ=0, h-1 y instantaneous weight fraction of unconverted feed λ dimensionless decay group (represents the “length” of decay), αtc α decay velocity constant, h-1 tc catalyst residence time, h θ normalized time-on-stream, t/tc



The decay parameters are introduced to account for the catalyst deactivation taking place as carbon deposits accumulate on the catalyst as feed conversion progresses. With the knowledge of the distribution of the catalyst in the riser reactor, the kinetics of gas oil conversion and the catalyst deactivation, it is possible to formulate the equations describing the feed conversion along the riser length. The governing equations for fixed bed and fluid bed reactors have been presented by Weekman27,28 and Weekman and Nace29. For a small diameter pilot plant riser reactor, it is a reasonable assumption to consider that the gas flow is in plug flow over the operating range of gas superficial velocities and catalyst fluxes of riser flow30. Hence, we can apply a unidimensional axial variation of concentration of reactant A (in our case the gas oil feed fraction boiling over 221°C). With this assumption and following Weekman27, the governing equation describing the feed conversion in a riser reactor takes the following form: KV dy = − Ac k o y 2 WHSV dz

(7)

where:

z axial reactor distance relative to total reactor length, z/Lr z axial distance in reactor, m Lr total riser length, m Ac : rate of decay of catalyst activity ko : rate constant for cracking oil feed, the same as in Eq. (5) KV = εg ρV/ρcat ρV : vapor density in riser, kg/m3 ρcat : catalyst density in riser, kg/m3 WHSV Weight Hourly Space Velocity, h-1 (Oil feed rate/per weight of catalyst holdup in reactor) 10 ACS Paragon Plus Environment

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Based on Eq. (7), Wollaston et al.31 developed equations, which are reproduced in Eqs. (8) and (9). These equations are considered a valuable tool for FCC applications, because they can easily be tested with either pilot or commercial FCC data.  X  −n   WHSV = Ac ∗ f1 ( feed ) ∗ f 2 (operating conditions ) t c 1 − X  

(Coke, wt %) ∗ WHSV = Ac

(8)

g1 ( feed ) g 2 (operating conditions) t c− n (9)

where: X feed volume fraction conversion, % KINC = X/(1-X) kinetic conversion, results from the integration of the second order kinetic Eq. (7) along the reactor length n decay exponent, 0.65 f1, f2 functions of feed quality and unit operating conditions g1, g2 functions of feed quality and unit operating conditions The effect of feed, catalyst and operating conditions on the f and g functions is determined with extensive pilot plant experimentation. The above analysis is based on the assumption that catalyst is quickly separated from oil vapors at the end of the riser. In our studies the disengagement zone temperature is maintained at 50°C lower than the riser exit temperature. This means that the catalytic and thermal reactions taking place at the disengagement zone are minimal. Using kinetic constants from literature32, it was found that the thermal cracking rates of representative compounds of the unconverted feed are about 10 times lower than the corresponding thermal rates in the riser. The effect of thermal decomposition in the riser is included in the reported results. In order to calculate product selectivity, the generalized kinetic models for fixed, fluid bed and riser reactors developed by Weekman27,28 and Weekman and Nace29 initially consider 11 ACS Paragon Plus Environment


three pseudocomponents: gasoline, gases and coke and unconverted feed. Many researchers33 have extended the number of pseudocomponents, as the analytical tools found wider use with the refiners. Dasila et al.34 presented the successful application of a 10-lump model in predicting product yields in a commercial FCC unit. The main contribution of their work is to connect the 10-lumps with detailed feed analysis, that is describing the feed composition with six hydrocarbon groups: light and heavy paraffins, light and heavy naphthenes, light and heavy aromatics. These six lumps together with 4-product lumps (unconverted feed, gasoline, gases and coke) participate in 25 cracking reactions. In this work a satisfactory agreement was obtained with commercial data although the hydrodynamics of the riser was simplified assuming among others that there is no slip between solid catalyst and vapor. Dewachtere et al.35 have presented a fundamental approach in calculating the rate of cracking of hydrocarbon species contained in a gas oil feedstock. This approach recognizes the fact that acid catalyzed hydrocarbon reactions, such as zeolite catalyzed processes, proceed via elementary carbenium ion reaction steps. It is based on the detailed knowledge of the mechanism of the various reactions involving carbenium ions. However, the application of the method to industrial feeds is difficult due to analytical complexity and computational limitations. Moreover, hydrogen transfer, coke formation, and coke fouling are not considered.

EXPERIMENTAL Nozzle design studies: For the nozzle design, the atomization of gas oil feed with steam at actual operating conditions was simulated with water and air at room temperature. Atomization depends on feed viscosity and surface tension. Although it is recognized that the properties of a


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vacuum gas oil at FCC process conditions are different from water, it has been found in our laboratory that relative differences among different nozzle designs can be established with waterair atomization at room temperature. A testing ring was used to test different nozzle designs. The atomization assembly consisted of the following parts: (a) A pre atomizer to premix the liquid with a gaseous stream. (b) The nozzle body consisting of a small diameter tube with length Lt and diameter Dt. (c) A small orifice tip with length Lo and diameter Do. The liquid flow rate was measured with a mass flowmeter, while air flow was measured using a variable area flowmeter. The quality of the produced spray was examined by measuring the droplet size with a Phase-Doppler Analyzer (PDA, Aerometrics). Different nozzle designs were tested. It was found that the droplet diameter depends on the nozzle length to diameter ratio (Lt/Dt), the dimensions of the diameter on the conical tip of the nozzle and the relative liquid to gas ratio. With the selection of proper design parameters and relative liquid to gas rates, it was found that it was possible to produce a spray with an average droplet diameter of 60 µm. Design of the riser entrance zone: The dimensions of the entrance zone were selected following a study in a cold flow model. Previous studies36 have established the importance of a cold flow model for the design of a Circulating Fluid Bed (CFB) system for oil shale retorting. Following similar techniques as described by Vasalos et al.36 using FCC catalyst (average particle size 60 µm), the geometry of the mixing zone was designed. The basis of the mixing zone design is to maintain a vapor superficial velocity less than 0.5 m/s. Thus, a critical catalyst density is maintained in order to provide good feed distribution over the catalyst particles and quick vaporization. The design of the mixing zone considered in the cold flow



experiments is depicted schematically in Figure 1 (riser entrance zone). A gas flow stream is split into two parts: one stream enters the reactor through the injection nozzle and a second one enters at the bottom through a distributor plate thus fluidizing the catalyst particles entering from the regenerator standpipe. The gas flow stream in the feed nozzle is required in order to secure the feed distribution in small droplets. The gas stream in the bottom of the mixing zone is necessary to fluidize the catalyst and thus secure quick momentum dissipation of the gas-feed stream exiting the tip of the atomizing nozzle. For the calculation of the solids density for the riser entrance zone two methods were followed: (a)

From pressure drop measurements according to the equation: ∆P = ρΒ g ∆H

(10)

where ρΒ is the bed density, ∆H is the distance between two pressure taps, and g is the acceleration of gravity. (b)

From proprietary correlations relating catalyst density with gas superficial velocity and

catalyst circulation rate. The catalyst density was calculated by dividing the catalyst weight in the entrance zone by its volume. The catalyst weight for each case was established by measurement of the solids holdup in cold flow studies. This was achieved by rapidly shutting off pneumatically actuated valves placed in the entrance of the gas and solids flows to the unit and the exit of the entrance zone. After the mixing zone was disassembled, the solids were collected and the amount of cracking catalyst in the mixing zone was determined. More details of this methodology are described by Vasalos et al.36 As reported, the bed densities calculated by the two methods are in reasonable agreement.


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FCC Pilot Plant Unit: The CPERI short length riser (1.46 m) FCC pilot plant was designed and constructed in early 90’s based on the results from a cold flow model. The pilot plant was modified to include a long (9.95 m) riser in 2006. A schematic diagram of the FCC pilot plant is shown in Figure 2a whereas in Figure 2b the location of the riser pressure taps is displayed. Detailed specifications of CPERI pilot plant have been reported by Lappas et al.8. Ιn a previous publication by Bollas et al.7, data were reported for a pilot plant design with a riser length equal to 1.46 m (short riser). For clarification purposes the 9.95 m length (long riser) will be the basis for this paper. It is, therefore, useful to provide a comparison of both riser designs in this publication. The long riser assembly is made of the following parts: 1. The riser entrance zone designated as zone 1 and shown in Figure 1. 2. The conical section connecting designated as zone 2 and shown in Figure 1. 3. The vertical section of the riser designated as zone 3 and shown in Figure 2b. 4. The down comer section of the riser designated as zone 4 and shown in Figure 2b. The relative volumes of the four zones are as follows: zone 1/zone 2/zone 3/zone 4= 1/1/15/20. These data imply that the pilot plant can easily be modified to run in the down comer mode. A comparison of key operating parameters of the pilot plant with a commercial unit is shown in Table 1. As shown, the key parameter, catalyst/oil ratio in the pilot plant, is very close to the commercial unit. Operating procedure: The pilot plant is fully automated and operates in a continuous catalyst circulation mode with catalyst regeneration like a commercial FCCU. Regenerated catalyst and preheated feed mix in the entrance of the riser-reactor and after separation of the catalyst at the exit of the riser, the catalyst flows to a stripper, where it is stripped with steam (0.6 g/min) from



hydrocarbon vapors trapped in its pores. The catalyst is returned via a lift line to the regenerator to burn off the coke, and then circulates back to the riser-reactor. For a given catalyst and feedstock to be tested in the FCC pilot plant the following procedure is applied: (a) Four kgs of catalyst are charged to the regenerator unit pneumatically from a hopper placed on top of the regenerator vessel. During this action the standpipe slide valve of the regenerator is closed. The unit is heated up through the heaters at the required temperatures. (b) The feedstock is charged in the feed drum and kept at the required preheat temperature. The main nitrogen inlet streams and the nitrogen aerations are adjusted to the required values. (c) During the start-up procedure the circulation of the catalyst between riser and regenerator takes place and the regenerator air is set at the desired stream value. When the recirculation of the catalyst in the unit reaches steady state, the feed stream is set at the desired value, usually at a rate of 25 g/min, and injected into the unit. (d) Subsequently the stripping steam rate is set at a value of about 0.6 g/min and introduced into the unit. (e) Steady state is usually achieved in two hours. (f) The time lag period (line-out) before the experiment start is about two to three hours, during which the unit is left to its steady state and no considerable adjustments are allowed. All analytical equipment is calibrated during this period and the unit is ready to carry out the required samples of experiment. (g) The run time of each experiment is about two hours. With the completion of the first experiment a second can be performed following a procedure similar to the above described. For further tests with another catalyst and/or another feedstock the evacuation of the unit from the catalyst and the feedstock is required. This can be done after feedstock pump shut-off and catalyst coke burn-out in the regenerator.


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Properties of the feeds studied: Lappas et al.8 have presented the range of feedstocks processed in the CPERI FCC Pilot Plant. In Table 2 the feed properties for three feedstocks used in this study are reported. Using the feed properties and standard ASTM procedures, the Molecular Weight (MW) of each feed was calculated.37 The feed MW is needed, because it is used to calculate the molar expansion during the cracking, and this in turn together with the inerts, determines the superficial vapor velocity in the riser. This is an important parameter to estimate the vapor void fraction, and hence, the catalyst holdup in the reactor zone. The latter together with the feed rate determines the Weight Hourly Space Velocity (Feed Rate-g/h/Catalyst holdupg).

RESULTS AND DISCUSSION Data Analysis The data reported in this part of the study - with feedstocks A and B and an Equilibrium Catalyst CAT1 (TSA=180 m2/g, ZSA=88 m2/g, MSA=92 m2/g, UCS=24.25 Å) - were obtained with isothermal operation of the main riser section of the reactor. This was achieved by operating the adiabatic heaters in a way to balance the endothermic heat of reaction and the wall losses along the riser. Riser exit temperature varied in the range of 526 to 538°C. A typical temperature profile along the riser is shown in Figure 3. With constant reaction temperature, feed conversion varied by changing the catalyst to oil ratio (C/O) and the Weight Hourly Space Velocity (WHSV). The following range of operating conditions was applied: Feed rate=15 to 25 g/min, Catalyst circulation=110 to 350 g/min, Feed preheat=170°C to 205°C. The catalyst circulation rate is calculated by dividing the rate of coke



production, as measured by the flue gas analysis of the regeneration gases, with the difference in carbon content of the spent and regenerated catalyst. For the calculation of the WHSV, in addition to the feed rate, the catalyst inventory in the reactor is needed. The latter is calculated by summing up the catalyst holdup in the four zones of the riser section. For each zone the catalyst inventory was calculated multiplying the volume of each zone with the local catalyst density. For the calculation of the catalyst density the gas voidage, εg, for each of the four zones needs to be calculated. For the specific feed and inert rates used in this study, the gas superficial velocity in each zone fell in the following range: Zone 1:

0.11 to 0.17 m/s

Zone 2:

0.46 to 0.75 m/s

Zone 3:

3.4 to 5.5 m/s

Zone 4:

1.2 to 2 m/s

As discussed earlier, the gas voidage in the entrance zone was established from correlations obtained in experiments in a cold flow unit with dimensions identical to the riser entrance of the pilot plant. For the calculation of the catalyst inventory in the riser section two approaches were followed: (a) via pressure drop measurements as indicated by Eq. (4) and (b) via calculation of the slip factor with Eq. (3). For approach (a) the pressure drop along the riser should mainly depend on the catalyst circulation rate. Indeed, as shown in Figure 4, the differential pressure along the riser monotonically increases with catalyst circulation rate. Figure 4 validates Eq. (4), which means that the static head due to catalyst inventory in the riser is the main component of the measured pressure drop. Regarding approach (b), for the calculation of the slip factor ψ, Eq.


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(3) was applied. Using Eq. (3) the slip factor was calculated as a function of gas superficial velocity. It was found that the average estimated value is close to 2, as reported by Matsen.22 Based on the above calculation procedures, it was found that the catalyst inventory in each zone fell in the following range: Zone 1:

10.5 to 11.5 grms

Zone 2:

6.2 to 7.2 grms

Zone 3:

6.2 to 14.3 grms

Zone 4:

8 to 18.2 grms

The narrow range for zones 1 and 2 is explained by the fact that the catalyst in zone 1 is fluidized in the bubbling bed and for zone 2 in the turbulent fluid bed regime. For these two zones the gas voidage is influenced only by gas superficial velocity with a low dependence on catalyst circulation rate. For zones 3 and 4, operating in the dilute phase mode, the gas voidage depends on catalyst circulation rate as reported by Zhu et al.38 and Wang et al.39. The WHSVs calculated with the reactor catalyst inventory from methods (a) and (b) are compared in Figure 5. It is shown that the catalyst inventory calculated from pressure drop measurements and from the slip factor approach, results in WHSVs which are in reasonable agreement. Pilot plant results – Feed conversion and coke yield Experimental data from feedstock A and an Equilibrium Catalyst (CAT1) were first used to calculate feedstock conversion and coke yield with a second order kinetics7,26,31. For this purpose pilot plant data were used at a constant reactor temperature of 526°C, and a feed rate ranging from 15 to 25 g/min.



The feed conversion calculated from experimental data was correlated with catalyst residence time in the reaction zone as reported by Bollas et al.7. It is shown that at constant reaction temperature and for a given feedstock and catalyst, feed conversion can be estimated from the catalyst residence time in the reaction zone. The second order reaction kinetics for feed conversion is verified with the data in Figure 6. From Bollas et al.7 it is also expected that the coke yield has a similar dependence on catalyst residence time as the feed conversion. This is shown in Figure 7. In Figures 6 and 7 a kinetic average temperature was used. This is an average temperature, which is calculated from the riser outlet temperature and the temperature profile along the riser. In order to test this hypothesis Eq. (8) was applied for the different sections of the riser, using the measured temperatures in the following reactor zones: (a) Zone 1: Nozzle exit to bottom of conical section (Figure 1), (b) Zone 2: Conical section, (c) Zone 3: Vertical section of riser, (d) Zone 4: Downflow section of riser. The conversion at the end of each zone was calculated by applying Eq. (8) to each section and to the total reactor length. The WHSV was calculated by using the catalyst inventory estimated by pressure drop data for each zone. Applying this procedure, the effect of reactor length (catalyst residence time) is shown in Figure 8. It is clearly shown that the estimated kinetic conversion at the end of zone 4 (end of riser) is close to the same value calculated by using a kinetic average temperature for the entire riser. The influence of riser length of an FCC pilot plant on catalyst residence time and product selectivity In an earlier publication by Bollas et al.7 we have reported feed conversion and coke yield correlations as a function of catalyst residence time for a pilot plant having a riser length of 1.46


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m. These data were compared with the results reported in this paper, where the riser length is 9.95 m. In both pilot plants the riser entrance zone is identical. In order to establish a direct comparison between the two pilot plants feedstock A and catalyst CAT1 were run at the same temperature, 526°C. The feed conversion varied by changing the catalyst to oil ratio. As shown in Figure 9 in order to reach high conversion, the short riser necessitates the use of much higher catalyst to oil (C/O) ratios compared with the long riser. As a result the catalyst residence time has greatly reduced. Using feed A and catalyst CAT1, the effect of catalyst residence time on conversion and coke yield is shown in Figures 10 and 11. These two figures clearly demonstrate that the second order kinetic model reliably predicts the effect of riser length on two key factors in Fluid Catalytic Cracking: Feed conversion and coke yield. As discussed earlier, the riser temperature was kept constant. Let us now consider how some of the product yields vary with catalyst to oil ratio in both the short and long riser pilot plants. Figures 12 and 13 show the variation in gasoline and LPG, respectively. The trends are consistent with Alvarez-Castro et al.5 data. Pilot plant validation – Commercial Data The next task is to compare pilot plant with commercial data. Two examples are discussed below: (a)

Prediction of catalyst activity and selectivity: In Figure 14 we show the variation of

feed conversion with catalyst to oil ratio. Experimental pilot plant data are compared with the same feed and catalyst as in a commercial unit. It is shown that by running an equilibrium Zeolite Y-ECAT catalyst, the catalyst activity measured from pilot plant matches commercial unit data. In addition, as shown in Figure 15, the gasoline yield measured for the pilot plant tests for the Y-ECAT is again matched with the commercial yield.



(b)

Prediction of slurry oil yield: A major issue today in the refining industry, is how to

minimize the amount of unconverted heavy (bottoms) product, which is value as bunker fuel. For this reason two new catalysts (CAT2 and CAT3) were evaluated in the pilot plant. Catalyst 2 has a high Zeolite/Matrix and catalyst 3 has a low Zeolite/Matrix surface ratio. The feedstock used (feedstock C) was a high sulfur feed with a specific gravity at 15°C of 0.928. The reactor temperature used was set at 538°C. The feedstock conversion changed from 67 to 77 wt% by varying the C/O from 5.5 to 11 and the Weight Hourly Space Velocity from 40 to 75. Detailed results from this study were presented by Joyal et al.3 in 2008. Bottom products from the two catalysts in the pilot plant are shown in Figure 16 and in a commercial unit3 are shown in Figure 17. It is clearly shown that the riser pilot plant unit tracks the performance of the catalysts in a commercial unit. The data imply that the selection of a catalyst from pilot plant experimentation is correctly scaled up to commercial application.

CONCLUSIONS This paper discusses a comparison of two pilot plants with different riser lengths (1.46 and 9.95 m). The pilot plants operated with a vacuum gas oil feed and an equilibrium zeolite catalyst at constant isothermal riser conditions. It was thus made possible to have a back to back comparison of riser length on catalyst to oil ratio to achieve the same feedstock conversion. Using existing hydrodynamic correlations, the distribution of the catalyst density along the riser was calculated. With the catalyst density, the reactor volume and the catalyst circulation rate, the catalyst residence time was estimated for a range of operating conditions. Using a simplified second order kinetic model, the effect of catalyst residence time on feed conversion and coke yield was established. For a given quality of feed and catalyst, at constant reactor temperature,


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wide variations in selectivity were observed with catalyst residence time, catalyst to oil ratio and vapor residence time. Pilot plant results were validated with commercial data. AUTHOR INFORMATION Corresponding Author: *Tel.: +30 2310498305, Fax: +30 2310498380, E-mail address: [email protected]



REFERENCES (1) Corma, A.; Sauvanaud, L. FCC Testing at Bench Scale: New Units, New Processes, New Feeds. Catal. Today 2013, 218-219, 107. (2) Bryden, K.; Weatherbee, G.; Habib, E.T. Flexible Pilot Plant Technology for Evaluation of Unconventional Feedstocks and Processes. Grace Catalysts Technologies Catalagram, Issue No. 113, 2013. (3) Joyal, C.L.M.; Westby, M.J.; Lappas, A.A.; Iatridis, D.K.; Vasalos, I.A. Pilot Plant Evaluation of FCCU Catalyst Technology and Use of Data for Commercial Catalyst Applications. Presented at the NPRA Annual Meeting, San Diego, March 2008; Paper AM-08-52. (4) Vasalos, I.A.; Lappas, A.A.; Iatridis, D.K.; Voutetakis, S.S. Design, Construction and Experimental Results of a Circulating Fluid Bed FCC Pilot Plant. In Proceedings of Circulating Fluidized Bed Technology V; Kwauk, V.M., Li, J., Eds.; Science Press: Beijing, China, 1996; pp 408-413. (5) Alvarez-Castro, H.C.; Matos, E.M.; Mori, M.; Martignoni, W.; Ocone, R. Analysis of Process Variables via CFD to Evaluate the Performance of a FCC Riser. Int. J. Chem. Eng. 2015, Volume 2015, Article ID 259603, 13 pages. (6) Wu, F.; Weng, H.; Luo, S. Study on Lumped Kinetic Model for FDFCC I. Establishment of Model. China Pet. Process. Pe. 2008, 2, 45. (7) Bollas, G.M.; Vasalos, I.A.; Lappas, A.A.; Iatridis, D.K. Modelling Small-Diameter FCC Riser Reactors. A Hydrodynamic and Kinetic Approach. Ind. Eng. Chem. Res. 2002, 41, 5410.


Page 24 of 40

Page 25 of 40 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


(8) Lappas, A.A.; Iatridis, D.K.; Papapetrou, M.C.; Kopalidou, E.P.; Vasalos, I.A. Feedstock and Catalyst Effects in Fluid Catalytic Cracking – Comparative Yields in Bench Scale and Pilot Plant Reactors. Chem. Eng. J. 2015, 278, 140. (9) Chen, Ye-Mon. Recent Advances in FCC Technology. Powder Technol. 2006, 163, 2. (10) Kumar, R.; Prasad, K.S.L. Studies on Pneumatic Atomization. Ind. Eng. Chem. Proc. Des. Dev. 1971, 10, 357. (11) Sovani, S.D.; Sojka, P.E.; Lefebvre, A.H. Effervescent Atomization. Prog. Energ. Combust. 2001, 27, 483. (12) Jedelsky, J.; Jicha, M.; Slama, J.; Otahal, J. Development of an Effervescent Atomizer for Industrial Burners. Energy Fuels 2009, 23, 6121. (13) Jedelsky, J.; Jicha, M. Spray Characteristics and Liquid Distribution of Multi-hole Effervescent Atomisers for Industrial Burners. Appl. Therm. Eng. 2016, 96, 286. (14) Buchanan, J.S. Analysis of Heating and Vaporization of Feed Droplets in Fluidized Catalytic Cracking Risers. Ind. Eng. Chem. Res. 1994, 33, 3104. (15) Nguyen, T.T.B.; Mitra, S.; Pareek, V.; Joshi, J.B.; Evans, G. Comparison of Vaporization Models for Feed Droplet in Fluid Catalytic Cracking Risers. Chem. Eng. Res. Des. 2015, 101, 82. (16) Mirgain, C.; Briens, C.; Del Pozo, M.; Loutaty, R.; Bergougnou, M. Modeling of Feed Vaporization in Fluid Catalytic Cracking. Ind. Eng. Chem. Res. 2000, 39, 4392. (17) Behie, L.A.; Bergougnou, M.A.; Baker, C.G.J.; Bulani, W. Jet Momentum Dissipation at a Grid of a Large Gas Fluidized Bed. Can. J. Chem. Eng. 1970, 48, 158. (18) Behie, L.A.; Bergougnou, M.A.; Baker, C.G.J.; Base, T.E. Further Studies on Momentum Dissipation of Grid Jets in a Gas Fluidized Bed. Can. J. Chem. Eng. 1971, 49, 557.



(19) Chenglin, E.; Fan, Y.; Zhang, K.; Zhang, H. Concentration Profile of Jet Gas in the Feed Injection Zone of a FCC Riser. Prog. Nat. Sci. 2008, 18, 1285. (20) Avidan, A.A.; Yerushalmi, J. Bed Expansion in High Velocity Fluidization. Powder Technol. 1982, 32, 223. (21) Richardson, J.F.; Zaki, W.N. The Sedimentation of a Suspension of Uniform Spheres under Conditions of Viscous Flow. Chem. Eng. Sci. 1954, 3, 65. (22) Matsen, J.M. Some Characteristics of Large Solids Circulation Systems. In Fluidization Technology vol. 2; Keairns, D.L., Ed.; Hemisphere Publishing Corp.: New York, 1976; pp 135-149. (23) Patience, G.S.; Chaouki, J.; Berruti, F.; Wong, R. Scaling Considerations for Circulating Fluidized Bed Risers. Powder Technol. 1992, 72, 31. (24) Pugsley, T.S.; Berruti, F. A Predictive Hydrodynamic Model for Circulating Fluidized Bed Risers. Powder Technol. 1996, 89, 57. (25) Das, M.; Meikap, B.C.; Saha, R.K. Voidage and Pressure Profile Characteristics of SandIron Ore-Coal-FCC Single-Particle Systems in the Riser of a Pilot Plant Circulating Fluidized Bed. Ind. Eng. Chem. Res. 2008, 47, 4018. (26) Blanding, F.H. Reaction Rates in Catalytic Cracking of Petroleum. Ind. Eng. Chem. 1953, 45, 1186. (27) Weekman, V.W. A Model of Catalytic Cracking Conversion in Fixed, Moving and FluidBed Reactors. Ind. Eng. Chem. Proc. Des. Dev. 1968, 7, 90. (28) Weekman, V.W. Kinetics and Dynamics of Catalytic Cracking Selectivity in Fixed-Bed Reactors. Ind. Eng. Chem. Proc. Des. Dev. 1969, 8, 385.


Page 26 of 40

Page 27 of 40 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


(29) Weekman, V.W.; Nace, D.M. Kinetics of Catalytic Cracking Selectivity in Fixed, Moving and Fluid-Bed Reactors. AIChe J. 1970, 16, 397. (30) Mahmoudi, S.; Seville, J.P.K.; Baeyens, J. The Residence Time Distribution and Mixing of the Gas Phase in the Riser of a Circulating Fluidized Bed. Powder Technol. 2010, 203, 322. (31) Wollaston, E.G.; Haflin, W.J.; Ford, W.D.; D’Souza, G.J. What Influences Cat Cracking. Hydrocarb. Process. 1975, 54, 93. (32) Behar, F.; Lorant, F.; Mazeas, L. Elaboration of a New Compositional Kinetic Schema for Oil Cracking. Org. Geochem. 2008, 39, 764. (33) Bollas, G.M.; Lappas, A.A.; Iatridis, D.K.; Vasalos, I.A. Five-lump Kinetic Model with Selective Catalyst Deactivation for the Prediction of the Product Selectivity in the Fluid Catalytic Cracking Process. Catal. Today 2007, 127, 31. (34) Dasila, P.K.; Choudhury, I.R.; Singh, S.; Rajagopal, S.; Chopra, S.J.; Saraf, D.N. Simulation of an Industrial Fluid Catalytic Cracking Riser Reactor Using a Novel 10Lump Kinetic Model and Some Parametric Sensitivity Studies. Ind. Eng. Chem. Res. 2014, 53, 19660. (35) Dewachtere, N.V.; Froment, G.F.; Vasalos, I.; Markatos, N.; Skandalis, N. Advanced Modeling of Riser-Type Catalytic Cracking Reactors. Appl. Therm. Eng. 1997, 17, 837. (36) Vasalos, I.A.; Tatterson, D.F.; Furlong, M.W.; Kowalski, T.L.; So, B.Y.C. Application of a Cold Flow Model in Testing the Feasibility of an Oil Shale Retorting Process. Ind. Eng. Chem. Res. 1991, 30, 1200.



(37) ASTM D2502-14, Standard Test Method for Estimation of Mean Relative Molecular Mass of Petroleum Oils from Viscosity Measurements, ASTM International, West Conshohocken, PA, 2014, www.astm.org. (38) Zhu, H.; Zhu, J. Characterization of Fluidization Behavior in the Bottom Region of CFB Risers. Chem. Eng. J. 2008, 141, 169. (39) Wang, C.; Li, C.; Zhu, J.; Wang, C.; Barghi, S.; Zhu, J. A Comparison of Flow Development in High Density Gas-Solids Circulating Fluidized Bed Downer and Riser Reactors. AIChE J. 2015, 61, 1172.


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Table 1. FCC Pilot Plant to Commercial FCCU Comparison

PARAMETERS Feed Rate (BPD) Feed Preheat Temp (°C) Riser Outlet Temp (°C) Cat to Oil for lab catalyst Cat to Oil for commercial catalyst

Ratio of Commercial to Pilot Plant 476190

0.67

Pilot Plant

Commercial

0.126 100 - 305 515 - 550 4 - 12 5 - 16

60000 190 - 290 515 - 538 4-9

Table 2. Feed properties FEED PROPERTIES Total nitrogen, ppmwt API Density15°C, g/L Micro Carbon Residue, %wt

FEED A 1133 19.6 936 0.29

FEED B 622 26.1 897 0.02

FEED C 1430 20.8 928 1.23

mass%_IBP mass%_10 mass%_20 mass%_30 mass%_50 mass%_70 mass%_80 mass%_90 mass%_FBP

1.5 23880 °C 327.4 393.5 414.4 429.3 454.4 483.1 500.1 524.2 551.6

1.5 2270 °C 200.7 296.1 331.7 355.2 395.7 435.6 459.5 494.5 539.3

15170 °C 314.7 383.2 407.0 423.7 453.3 488.2 509.4 540.6 561.2

Molecular Weight

334.3

288.5

338.1

Refractive_Index_70°C Sulfur, ppmwt



Figure captions

Figure 1. Mixing zone geometry Figure 2a. Schematic of the Riser Pilot Plant Figure 2b. Location of the riser pressure taps Figure 3. Typical temperature profile along the riser – FEED A, CAT1 Figure 4. Pressure drop along riser vs catalyst circulation rate – FEEDS A & B, CAT1, Riser Length = 9.95 m Figure 5. Recalculated WHSV riser vs WHSV riser from Riser ∆P – FEEDS A & B, CAT1, Riser Length = 9.95 m Figure 6. LN(WHSV*KIN CONV) vs LN(tc) – Riser Length = 9.95 m Figure 7. LN(WHSV*COKE) vs LN(tc) – Riser Length = 9.95 m Figure 8. Effect of riser geometry on conversion Figure 9. Conversion vs C/O ratio Figure 10. LN(WHSV*KIN CONV) vs LN(tc) Figure 11. LN(WHSV*COKE) vs LN(tc) Figure 12. Gasoline vs C/O ratio Figure 13. LPG vs C/O ratio Figure 14. Comparison of pilot plant with commercial data - Conversion vs C/O Figure 15. Comparison of pilot plant with commercial data - Gasoline vs conversion Figure 16. Comparison of slurry yields with FEED C and CAT2 & CAT3 - Pilot plant Figure 17. Comparison of slurry yields with FEED C and CAT2 & CAT3 - Commercial data


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Figure 1



Figure 2a

Figure 2b

Figure 3 32 ACS Paragon Plus Environment

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Figure 4

Figure 5



Figure 6

Figure 7


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Figure 8

Figure 9



Figure 10

Figure 11


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Figure 12

Figure 13



Figure 14

Figure 15


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Figure 16

Figure 17



FOR TABLE OF CONTENTS ONLY


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Influence of Riser Length of a Fluid Catalytic Cracking Pilot Plant on

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